Processes and apparatus for continuous solution polymerization

ABSTRACT

The invention relates to processes and plants for continuous solution polymerization. Such plant and processes include a pressure source, a polymerization reactor, downstream of said pressure source, pressure let-down device, downstream of said polymerization reactor, and a separator, downstream of said pressure let-down device, wherein said pressure source is sufficient to provide pressure to said reaction mixture during operation of said process plant to produce a single-phase liquid reaction mixture in said reactor and a two-phase liquid-liquid reaction mixture in said separator in the absence of an additional pressure source between said reactor and said separator.

FIELD OF INVENTION

[0001] The invention relates to processes and apparatus for continuoussolution polymerization. The invention relates especially to suchprocesses and apparatus using single site soluble transition metalcatalysts, in particular those known as metallocene catalysts. Theinvention furthermore relates especially to such processes and apparatusthat provide for improved control of hydrogen level, and more especiallyfor improving series reactor operation using single site solubletransition metal catalysts.

BACKGROUND OF INVENTION

[0002] Continuous solution polymerization processes generally involvethe addition of catalyst to a monomer and solvent mixture. The mixturemay be back-mixed giving a uniform polymer in an environment withsubstantially no concentration gradients. WO 94/00500 (Pannell, et al.)describes a solution polymerization using metallocene in a continuousstirred tank reactor, which may be in a series reactor arrangement tomake a variety of products.

[0003] For the purposes of this patent specification the term“metallocene” is herein defined to contain one or more cyclopentadienylmoiety in combination with a transition metal of the Periodic Table ofElements.

[0004] The heat of the polymerization reaction can be absorbed by thepolymerization mixture, causing an exotherm. Alternatively, or inaddition, the heat of reaction can be removed by a cooling system, byexternal cooling of the walls of the reactor vessel, or by internallyarranged heat exchange surfaces cooled by a heat exchange fluid.

[0005] In the course of the polymerization, typically, a predominantamount (over 50 mol %) of the monomer is consumed and the polymer formedis dissolved in the solvent. The higher the concentration of thepolymer, the higher the viscosity of the polymerization reaction mixturecontaining the polymer, solvent, and unreacted components. The mixturepasses from the polymerization reactor to a finishing section in whichpolymer, solvent and unreacted monomer are separated. In the course offinishing, solvent and unreacted monomer are progressively removed fromthe polymerization mixture until the polymer can be formed into a solidpellet or bale. The separated solvent and monomer can be recycled to thepolymerization reactor.

[0006] It is well known from extensive literature sources that polymersolutions can undergo phase separation at the lower critical solutiontemperature, with phase separation being encouraged by highertemperatures and/or by lower pressures. Solvents selection alsoinfluences the conditions where phase separation occurs.

[0007] The phenomenon of phase separation is firstly a consideration inthe selection of the polymerization solvent. Appropriate polymerizationmonomer conversions, especially of the volatile monomers, temperatures,and pressures have to be selected for given polymer/solvent combinationconditions to avoid unwanted phase separation inside the reactor.Solvents such as hexane may require an elevated pressure in excess of 50bar to avoid two-phase conditions for olefin polymerization; solventssuch as octane can maintain homogeneous one-phase conditions at lowerpressures.

[0008] The phenomenon of phase separation can secondly be exploitedafter the reaction step to separate volatile solvent and unreactedmonomer components on the one hand, and polymer on the other hand. Inthat case, separation at temperature well above the lower criticalsolution temperature is encouraged to allow the polymer to form aconcentrated phase. Some earlier articles explain the general principlesof which we are aware are: “A low-energy Solvent Separation Method,” byT. G. Gutowski et al, Polymer Engineering; “Solvents” by C. A. Irani etal. in Journal of Applied Polymer Science Vol 31, 1879-1899 (1986);“Separating Polymer Solutions with Supercritical Fluids,” by Mark A.McHugh et al in Macromolecules 1985, 18, 674-680; “Critical dynamics andphase separation kinetics etc,” by Hajime Tanaka in Journal of ChemicalPhysics 100 (7) Apr. 1, 1994 p 5323-5337; “Short Chain Branching Effecton the Cloud Point Pressures of Ethylene Copolymers etc.,” by S. J. Hanet al. in Macromolecules 1998, 31, 2533-2538.

[0009] U.S. Pat. No. 3,726,843 described a process for making EPDM.Liquid phase separation has also been exploited to remove solvent fromthe polymerized mixture exiting from the polymerization reactor inMitsui EP-552945-A (U.S. Pat. No. 5,599,885), which shows a continuoussolution polymerization process with a metallocene catalyst. Hydrogen isadded in the examples to avoid higher molecular weights at the lowoperating temperature. The pressure and temperature are raised to permita subsequent pressure drop, that leads to the formation of separate leanand concentrated phases. Catalyst emerging from the reactor is recycled.

[0010] EP-552945-A does not disclose that the polymerization process maybe conducted at elevated pressures to provide a wide range of polymersand outputs in the same plant arrangement. EP-552945-A uses anauto-refrigerated reactor in which the solvent is allowed to boil whichfavors low pressure operation. EP-552945-A does not suggest exploitingthe initial elevated pressure in the finishing section.

[0011] While the single site, metallocene catalysts have a highactivity; that activity is often sustained under conditions in whichphase separation would occur at elevated temperatures. Continuedpolymerization activity during phase separation may influence polymercharacteristics undesirably.

[0012] The use of single site catalysts is associated with poorsolubility in the aliphatic hydrocarbon, saturated, non-polar solventsused for homogeneous solution polymerization. As a result, an aromaticcatalyst solvent, such as toluene may have to be used. This in turn cancomplicate solvent separation to prevent toluene build up in thereactor, and lead to environmental pollution and added maintenanceexpenditure. EP-552945-A tries to avoid the use of toluene by slurryingthe catalyst, comprising alumoxane as activator, in the polymerizationsolvent.

[0013] In some solution processes (see WO 98/02471 Kolthammer) thepolymerized mixture is flashed off in two stages, whereby the solventand unreacted monomer are converted to a vapor phase. Efficientextraction of solvent, etc., requires low vapor pressures and vaporphase compression or condensation followed by pumping for subsequentseparation stages. Pumping is used to convey polymer from flashseparation stages to a final devolatilizing extruder

[0014] U.S. Pat. No. 3,912,698 uses a heat exchanger for a liquidrecycle stream to permit an increase in reactor capacity while reducingfouling in the context of a multiple flash to remove volatiles.

[0015] The use of single site catalysts is also associated with thegeneration of hydrogen through beta-hydride abstraction. Such hydrogen,when recycled back to reactor feed, can act as a modifier to reduce themolecular weight of the polymer. The amount of hydrogen established inpolymerization may have to be increased or decreased depending on thetarget molecular weight.

[0016] In solution plants, solvent selection, operating temperatures,and purification systems have to be designed for a particular operatingwindow for the desired polymerization process. Metallocene catalystspermit a wide variety of polymers to be made in terms of comonomercontent, molecular weight, etc. Optimum production performance for agiven type of polymer may be obtained with a particular metallocenewithin a specific operating window. Different types of polymer may thenhave to be produced in different plant lay-outs. There is, therefore, aneed for a plant design that can be used more flexibly for differenttypes of polymers and metallocene catalysts, and which also can beadapted more easily to evolving metallocene catalyst technologies thancurrent designs of solution polymerization plants.

[0017] There is also a need for a plant design that permits moreextensive molecular weight control through control of the hydrogenlevels. There is an special need for such control that is compatiblewith series reactor operation that permits well separatedsplit-operating conditions between the first and second reactor (onewhich permits feeding very low levels of hydrogen to one of the tworeactors, while feeding large amounts of hydrogen to the other reactor).

[0018] There remains a need for an improved continuous solution processand plant which provides one or more of the following benefits:producing polymer economically across a broad range of operating windowsincluding varying polymerization temperatures; producing a broadspectrum of polymers, particularly polymers of widely varying averagemolecular weights, molecular weight distributions, and/or comonomercontents; permitting the production of polymers having useful molecularweights at high temperatures (above 150° C.); accommodating a broadrange of catalyst performance; reducing energy consumption, especiallyin finishing, and reducing environmental discharge; and reducing oravoiding fouling in the recycle and purification systems while usinghighly active metallocene type catalysts with unreacted monomer andtemperature during separation processes.

[0019] It would be particularly useful to provide a process and plantwhich can adjust the process window to optimize performance for a givenpolymer type and catalyst, such that metallocene catalyst can be used toperform at a high activity within that window; while at the same timepermitting a broad range of optimized performance windows for differentpolymer types and catalysts. It would also be beneficial to provide aprocess and plant which could facilitate operating at such high catalystactivities in the same finishing equipment, which can be used in alargely closed system with substantial recycling of all non-polarsolvent and monomer components; with minimal contamination and minimalneed to eliminate polar impurities contained in such non-polar recycle,however derived (catalyst residue; scavenger, etc), using a simpleremoval technique, and without using a stripping agent such as waterwhich would contaminate the recycle.

[0020] For additional background, see also WO 94/00500 and WO92/14766.

SUMMARY OF INVENTION

[0021] The numbers shown in brackets refer to the numbering of itemsshown in the drawings and intended for illustration and facilitatingunderstanding and are not intended to limit the disclosure to the itemsillustrated.

[0022] The invention relates generally to a process for continuoussolution polymerization of a feed (2, 4, 58) of olefinically unsaturatedmonomers in a hydrocarbon solvent under pressure, having a continuousstirred tank reactor arrangement (8), to which a single site catalyst issupplied, to form a polymer containing polymerization reaction mixture,and downstream thereof a separating means for continuous separation ofthe solvent (14, 34, 40) and unreacted monomer from the mixture, whichseparating means (14, 34, 40) includes at least an initial liquid phaseseparator to separate the polymerization mixture into a lean phase (20)and a concentrated phase (22). The terms “lean” and “concentrated” or“polymer rich” refers to amount of polymer in the solvent. “Lean”indicates that the solvent contains no polymer or such low amounts ofpolymer so as to not interfere with subsequent recycling.

[0023] The catalyst is preferably a bulky ligand transition metalcatalyst. The bulky ligand contains a multiplicity of bonded atoms,preferably carbon atoms, forming a group, which may be cyclic with oneor more optional hetero-atoms. The bulky ligand may be metallocene-typecyclopentadienyl derivative, which can be mono- or poly-nuclear. One ormore bulky ligands may be bonded to the transition metal atom. The bulkyligand is assumed, according to prevailing scientific theory, to remainin position in the course of polymerization to provide a homogenouspolymerization effect. Other ligands may be bonded or coordinated to thetransition metal, preferably detachable by a cocatalyst or activator,such as a hydrocarbyl or halogen-leaving group. It is assumed thatdetachment of any such ligand leads to the creation of a coordinationsite at which the olefin monomer can be inserted into the polymer chain.The transition metal atom is a Group IV, V or VI transition metal of thePeriodic Table of Elements. The transition metal atom is preferably aGroup IVB atom.

[0024] In a first aspect of the invention a high capacity, low viscositypump (3) raises the pressure of the feed (2, 4, 58) to at least 75 barand causes the mixture to pass from the reactor (8) through a heatingarrangement (12, 16) up to a pressure reducing means (18) upstream ofthe liquid phase separator (14) through the action of the pump (3) andin the absence of further pumping means between the reactor (8) and thepressure reducing means (18). The term “high capacity, low viscositypump” refers generally to a pump with a capacity sufficient topressurize the whole of the feed at a viscosity not affected by thepresence of viscvosity increasing dissolved polymer molecules. Asuitable pump type is a centrifugal pump.

[0025] The heating arrangement may comprise a first heating stage (12)and a second heating stage (16). The first stage (12) is a heatintegrating heat exchanger designed to recover heat that would otherwisebe lost when the lean phase (20) was cooled in a final cooler (24). Thesecond stage (16) uses any appropriate heat utility of suitabletemperature to finish the heating step.

[0026] The reactor arrangement may be single reactor or a plurality,prefereably two, reactors arranged in series, or less preferably inparallel.

[0027] In this first aspect, acatalyst killer (10) is added downstreamof reactor arrangement (8) (in the case of series reactors that meansthat the killer is added downstream of the last polymerization reactor)and upstream of the heating arrangement (12, 16) and the liquid phaseseparator (14) to suppress further polymerization of the heatedpolymerization mixture undergoing separation, the lean phase (20) beingpassed through a cooling arrangement, which may comprise the heatintegrating exchanger (12) and a final cooler (24), and optionally adrier (32) back to the inlet side of the pump (3); the concentratedphase (22) being subjected to additional solvent removal downstream toobtain a solid polymer.

[0028] In a second aspect the lean phase (20) is passed in liquid formto a means (26) for removing hydrogen added to or generated duringpolymerization, which means (26) comprises a means for contacting astripping vapor with the lean phase in a countercurrent flow arrangementto concentrate the hydrogen in the vapour phase for removal from thelean phase recycle. This aspect may be practiced in processes and plantsnot employing the features of the first aspect.

DESCRIPTION OF THE DRAWINGS

[0029]FIG. 1 shows a schematic lay-out of a plant according to theinvention and a process flow according to the invention.

[0030]FIG. 2 shows a phase diagram illustrating the operation of aliquid phase separator used in the plant and process according to theinvention.

[0031]FIG. 3 shows the details of the lay-out of a hydrogen strippingarrangement for the plant of FIG. 1 to provide wide ranging molecularweight control.

[0032] The numbers in brackets below again refer to the correspondingfeatures in the drawings included for illustration and ease ofunderstanding. The invention includes other forms of the featuresindicated by the numbers in brackets in addition to those shown in thedrawings as would be apparent to person skilled in the art.

DETAILED DESCRIPTION OF GENERAL AND FIRST ASPECTS

[0033] By raising the pressure by the pump (3) to above 75 bar, theformation of two-phase conditions is avoided in the reactor arrangement(8) and heating arrangement (12, 16) under a wide range of temperatureand polymerization conditions. Hence a wide variety of metallocenecatalysts can be used in the process. Such a wide range of metallocenecatalysts can be used to make high and/or low average molecular weightmaterials under optimized production conditions. Use of a separate pumpto boost the pressure of the viscous polymerization mixture between thereactor and the liquid phase separator can be avoided, [such pumps aremuch more costly than the low viscosity feed pump (3)]. The pressure ofthe pump (3) also cascades through the process and combines with theabsence of vaporization for the initial solvent separating stage toreduce overall pumping needs during finishing. The pressure of the pump(3) advances the viscous polymerization mixture to the pressure reducingmeans (18) upstream of the liquid phase separator (14) without allowingphase separation prior to the pressure reducing means (18). In apreferred form of the invention, the pressure of the pump (3)additionally advances one or both separated phases to further downstreamfractionating systems or purification devices such as high pressureflash separation devices or low pressure flash separating devices.

[0034] The integral killer arrangement permits the temperature to beincreased without risking further polymerization so facilitating directrecycle, after removal of any surplus killer, of separated solvent andmonomer to the inlet side of the pump (3). With term “direct” is meantthat the lean phase does not need to be fractionated. The plant can beused under a wide variety of conditions to make a wide variety ofpolymers and is at the same time of simple construction.

[0035] Energy consumption per unit polymer produced is low, with simplesolvent recovery and energy integration systems [such as heatintegrating exchanger (12)], which can be employed to minimizedischarges to atmosphere and to recover heat from the effluent on theliquid phase separator (14).

[0036] Preferably the liquid phase separator (14) is connected to a lowpressure separator (34), arranged downstream, which receives theconcentrated phase from the liquid phase separator (14). Preferably alow-pressure separator (34) operates at a pressure sufficient to allowfeeding of the vapor phase to the fractionating and purification systemwithout requiring a separate compressor, and said pressure is generally2 bar guage or more. In order to accommodate production of polymers witha wide range of molecular weights, this pressure in the low pressureseparator (34) can be raised to a high level, between 3 and 20 barg, toadjust solution viscosity to facilitate feeding of the concentratedpolymer solution to the final devolatizing stage. Thus the volatilephase removed from a concentrated phase is conveyed simply to afractionating tower (36) as a vapor, arranged downstream of the lowpressure separator (34), for purification. In some prior artarrangements where solvents, etc., are drawn off under a low pressure inthe vapor phase, the extracted volatiles must be condensed and passedthrough pumping means for subsequent further separation steps.

[0037] Preferably, the process uses a non-polar solvent which does notcoordinate or interfere in a meaningful way so as to inhibit thecatalytic action of the catalyst system. Preferably the process uses alow boiling, alkane based solvent, optionally mixtures of alkanes, whichmay be linear or branched, such as those having from 4 to 10 carbonatoms, preferably in the range of 5-7 carbon atoms, optionally inadmixture with other alkanes of a higher or lower molecular weight. Thepolymer may be derived of monomers predominantly comprising mono-olefinssuch as ethylene or propylene or other higher alpha-olefins having from4 to 10 carbon atoms. This combination provides a mixture which can beeasily separated inside the liquid phase separator.

[0038] Considerable energy can be preserved by providing that thepolymerization mixture from the reactor (8) is heated to the temperaturebefore reaching the separator (14) successively by an upstream heatintegration exchanger (12) and a downstream trim heat exchanger (16) andby providing that the lean phase (20) from the separator (14) is used tosupply heat to the upstream one (12) of said heat exchangers.

[0039] Working pressures in the process of the invention can be 80 baror more, 90 bar or more; 95 bar or more and especially 120 bar or more,or even 140 bar or more. The upper pressure limit is not criticallyconstrained but typically can be 200 bar or less, preferably 140 bar orless, or 120 bar or less. The pressure should suffice to keep thereactor solution in a single phase up to the point of the pressurereducing means (18), and to provide the necessary working pressure toconvey the fluids through the plant.

[0040] The invention in another aspect also relates to a plant adaptedto perform the process described above which is suited to operate withinthe performance envelopes indicated below and with the SSC indicatedbelow. More aspects will be apparent from the claims.

[0041] The feed temperature may vary depending on the available exothermand extent of monomer conversion desired to reach the polymerizationtemperature. Advantageously the temperature is at least minus 40° C.,suitably, at least −20° C., 0° C., 20° C. or 40° C. in certaincircumstances. The polymerization temperature is constrained by themolecular weight desired, allowing for the influence of any hydrogenadded. In a series reactor process the temperature in the successivereactors can be raised progressively in increments depending on thenature of the polymerization taking place in such reactors.Advantageously, the polymerization temperature for polymers comprisingpredominantly ethylene derived units is at least 100° C., preferably atleast 150° C. or even (for lower molecular weight materials) 200° C. ormore. The temperature should not exceed the polymerization decompositiontemperature or the temperature at which the catalyst can sustain thepolymerization reaction.

[0042] Overall the exotherm may lead to a temperature differentialbetween the inlet temp of the polymerization reactor and the outlet offrom 50 to 220 or up to 250° C. By feeding at minus 40° C. and allowingthe exotherm to raise the temperature to 210° C., a highly efficientprocess may result for producing lower molecular weight polymers. Forhigher molecular weight polymers, the temperature rise may need to beconstrained via warmer feed and/or lower reactor temperatures to avoidexcessive viscosity in the reactor solution that would degrade reactormixing performance, thereby leading to non-uniform polymers.

[0043] Monomer concentration depends on the target polymer type andmolecular weight, the associated conversions of monomer to polymer andoperating temperature. Advantageously, the monomer partial pressureshould be 30% or more of the total partial pressure of volatilecomponents in the polymerization reactors; especially 40% or more, andshould preferably not exceed 80%, 70% or especially 60%. The totalpartial pressure of all components should be less than 100% of thereactor pressure to avoid formation of vapor bubbles. In general, highermonomer partial pressures are preferred to improve the liquid phaseseparation in the liquid phase separator (14).

[0044] In its broadest form, the invention can be performed with any SSC(single sited catalyst). These generally contain a transition metal ofGroups 3 to 10 of the Periodic Table; and at least one ancillary ligandthat remains bonded to the transition metal during polymerization.Preferably the transition metal is used in a cationic state andstabilized by a cocatalyst or activator. Especially preferred aremetallocenes of Group 4 of the Periodic table such as titanium, hafniumor zirconium which are used in polymerization in the d⁰ mono-valentcationic state and have one or two ancillary ligands as described inmore detail hereafter. The important features of such catalysts forcoordination polymerization are the ligand capable of abstraction andthat ligand into which the ethylene (olefinic) group can be inserted.

[0045] The metallocene can be used with a cocatalyst which may bealumoxane preferably methylalumoxane having an average degree ofoligomerization of from 4 to 30 as determined by vapor pressureosmometry. Alumoxane may be modified to provide solubility in linearalkanes or be used in a slurry but is generally used from a toluenesolution. Such solutions may include unreacted trialkyl aluminum and thealumoxane concentration is generally indicated as mol Al per liter,which figure includes any trialkyl aluminum which has not reacted toform an oligomer. The alumoxane, when used as cocatalyst, is generallyused in molar excess, at a mol ratio of 50 or more, preferably 100 ormore, and preferably 1000 or less, preferably 500 or less, relative tothe transition metal.

[0046] The SSC should preferably be selected from among a broad range,of available SSC's, to suit the type of polymer being made and theprocess window associated therewith in such a way that the polymer isproduced under the process conditions at an activity of at least 40,000g polymer per gram SSC (such as a metallocene), preferably at least60,000 or even in excess of 100,000 g polymer per g SSC. Thisspecification and examples exemplify some of the options. By enablingthe different polymers to be produced in different operating windowswith an optimized catalyst selection, the SSC and any ancillary catalystcomponents can be used in small quantities, with optionally also usingsmall amounts of scavengers. The killer can be used in equally smallamounts and the various cost-effective methods can then be introduced toallow the non-polar solvent to be recycled and subjected to treatment toremove polar contaminants before re-use in the polymerizationreactor(s).

[0047] The metallocene may be also be used with a cocatalyst which is anon- or weakly coordinated anion (the term non-coordinating anion asused herein includes weakly coordinated anions. The coordination shouldbe sufficiently weak in any event, as evidenced by the progress ofpolymerization, to permit the insertion of the unsaturated monomercomponent.) The non-coordinating anion may be supplied and reacted withthe metallocene in any of the manners described in the art.

[0048] The precursor for the non-coordinating anion may be used with ametallocene supplied in a reduced valency state. The precursor mayundergo a redox reaction. The precursor may be an ion pair of which theprecursor cation is neutralized and/or eliminated in some manner. Theprecursor cation may be an ammonium salt as in EP-277003 and EP-277004.The precursor cation may be a triphenylcarbonium derivative.

[0049] The non-coordinating anion can be a halogenated,tetra-aryl-substituted Group 10-14 non-carbon element-based anion,especially those that are have fluorine groups substituted for hydrogenatoms on the aryl groups, or on alkyl substituents on those aryl groups.

[0050] The effective Group 10-14 element cocatalyst complexes of theinvention are, in a preferable embodiment, derived from an ionic salt,comprising a 4-coordinate Group 10-14 element anionic complex, where A⁻can be represented as:

[(M)Q₁Q₂ . . . Q_(i)]⁻,

[0051] where M is one or more Group 10-14 metalloid or metal, preferablyboron or aluminum, and each Q is a ligand effective for providingelectronic or steric effects rendering [(M′)Q₁Q₂ . . . Q_(n)]⁻ suitableas a non-coordinating anion as that is understood in the art, or asufficient number of Q are such that [(M′)Q₁Q₂ . . . Q_(n)]⁻ as a wholeis an effective non-coordinating or weakly coordinating anion. ExemplaryQ substituents specifically include fluorinated aryl groups, preferablyperfluorinated aryl groups, and include substituted Q groups havingsubstituents additional to the fluorine substitution, such asfluorinated hydrocarbyl groups. Preferred fluorinated aryl groupsinclude phenyl, biphenyl, naphthyl and derivatives thereof.

[0052] The non-coordinating anion may be used in approximately equimolaramounts relative to the transition metal component, such as at least0.25, 15 preferably 0.5, and especially 0.8 and such as no more than 4,preferably 2 and especially 1.5.

[0053] Representative metallocene compounds can have the formula:

L^(A)L^(B)L^(C) _(i) MDE

[0054] where, L^(A) is a substituted cyclopentadienyl orhetero-cyclopentadienyl ancillary ligand π-bonded to M; L^(B) is amember of the class of ancillary ligands defined for L^(A), or is J, ahetero-atom ancillary ligand π-bonded to M; the L^(A) and L^(B) ligandsmay be covalently bridged together through a Group 14 element linkinggroup; L^(C) _(i) is an optional neutral, non-oxidizing ligand having adative bond to M (i equals 0 to 3); M is a Group 4 or 5 transitionmetal; and, D and E are independently mono-anionic labile ligands, eachhaving a σ-bond to M, optionally bridged to each other or L^(A) orL^(B.) The mono-anionic ligands are displaceable by a suitable activatorto permit insertion of a polymerizable monomer or macro-monomer caninsert for coordination polymerization on the vacant coordination siteof the transition metal component.

[0055] Representative non-metallocene transition metal compounds usableas SSC's also include tetrabenzyl zirconium, tetrabis(trimethylsiylmethyl) zirconium, oxotris(trimethlsilylmethyl)vanadium, tetrabenzyl hafnium, tetrabenzyl titanium, bis(hexamethyldisilazido)dimethyl titanium, tris(trimethyl silyl methyl) niobiumdichloride, and tris(trimethylsilylmethyl) tantalum dichloride.

[0056] Additional organometallic transition metal compounds suitable asolefin polymerization catalysts in accordance with the invention will beany of those Group 3-10 that can be converted by ligand abstraction intoa catalytically active cation and stabilized in that active electronicstate by a non-coordinating or weakly coordinating anion sufficientlylabile to be displaced by an olefinically unsaturated monomer such asethylene.

[0057] More preferred are metallocenes which are biscyclopentadienylderivatives of a Group IV transition metal, preferably zirconium orhafnium. See ExxonMobil WO9941294. These may advantageously bederivatives containing a fluorenyl ligand and a cyclopentadienyl ligandconnected by a single carbon and silicon atom. See ExxonMobil WO9945040;and WO9945041 and. Most preferably the Cp ring is unsubstituted and/orthe bridge contains alkyl substituents, suitably alkylsilyl substituentsto assist in the alkane solubility of the metallocene. See WO0024792 andWO0024793. Other possible metallocenes include those in WO01/58912

[0058] Dow EP418044 uses a monocyclopentadienyl compound similar thatthat EP416815. Similar compounds are described in ExxonMobil EP-420436.Sumitomo WO9703992 shows a catalyst in which a single Cp species and aphenol are linked by a C or Si linkage, such asMe2C(Cp)(3-tBu-5-Me-2-phenoxy)TiCl2. Nova WO200105849 disclosesCp-phosphinimine catalysts, such as (Cp)((tBu)3P═N—)TiCl2.

[0059] Other suitable metallocenes may be bisfluorenyl derivatives orunbridged indenyl derivatives which may be substituted at on eor morepositions on the fused ruing with moieties which have the effect ofincreasing the molecular weight and so indirectly permit polymerizationat higher temperatures such as described in EP693506 and EP780395.

[0060] When using the catalysts of the invention, the total catalystsystem will generally additionally comprise one or more organometalliccompounds as scavenger. Such compounds as used in this application ismeant to include those compounds effective for removing polar impuritiesfrom the reaction environment and for increasing catalyst activity.Impurities can be inadvertently introduced with any of thepolymerization reaction components, particularly with solvent, monomerand catalyst feed, and adversely affect catalyst activity and stability.It can result in decreasing or even elimination of catalytic activity,particularly when ionizing anion pre-cursors activate the catalystsystem. The impurities, or catalyst poisons include water, oxygen, polarorganic compounds, metal impurities, etc. Preferably steps are taken toremove these poisons before introduction of such into the reactionvessel, for example by chemical treatment or careful separationtechniques after or during the synthesis or preparation of the variouscomponents, but some minor amounts of organometallic compound will stillnormally be used in the polymerization process itself.

[0061] Typically these compounds will be organometallic compounds suchas the Group-13 organometallic compounds of U.S. Pat. Nos. 5,153,157,5,241,025 and WO-A-91/09882, WO-A-94/03506, WO-A-93/14132, and that ofWO 95/07941. Exemplary compounds include triethyl aluminum, triethylborane, tri-isobutyl aluminum, tri-n-octyl aluminum, methylalumoxane,and isobutyl alumoxane. Alumoxane also may be used in scavenging amountswith other means of activation, e.g., methylalumoxane andtri-isobutyl-aluminoxane with boron-based activators. The amount of suchcompounds to be used with catalyst compounds of the inventions isminimized during polymerization reactions to that amount effective toenhance activity (and with that amount necessary for activation of thecatalyst compounds if used in a dual role) since excess amounts may actas catalyst poisons.

[0062] The process and the plant used in the process are designed asexplained above to permit polymerization of a wide variety of polymertypes and molecular weights. Generally speaking the polymers are derivedfrom either ethylene or propylene as the dominant (more than 50 mol %)component. Polymers may preferably contain from 5 to 40 mol % ofcomonomer to vary crystallinity and flexibility. The comonomers may bealpha-olefins (under which term cyclic olefins such as styrene areincluded) having from 2 to 20 carbon atoms, such as ethylene (in thecase of the polymer consisting predominantly of propylene derived units)butene-1, hexene-1, octene-1. Amounts of dienes such as hexadiene, vinylnorbornene, ethylidene norbornene (ENB), norbornadiene etc may beincluded to promote unsaturation and/or the formation of longer branchesthemselves made from polymerized monomer derived units.

[0063] In the case of plastomer, the polymer which may be producedinclude the following aspects: Preferably the comonomer is analpha-olefin having from 3 to 15 carbon atoms, more preferably 4 to 12carbon atoms and most preferably 4 to 10 carbon atoms. Ethylene can bepolymerized with at least two comonomers to form a terpolymer. Monomeris generally polymerized in a proportion of 70.0-99.99, preferably 70-90and more preferably 80-95 or 90-95 mole % of ethylene with 0.01-30,preferably 3-30 and most preferably 5-20 mole % comonomer. For thepurposes of this patent specification the molecular weight distributionof a polymer can be determined with a Waters Gel PermeationChromatograph equipped with Ultra-styrogel 5 columns and a refractiveindex detector. The operating temperature of the instrument was set at145° C., the eluting solvent was trichlorobenzene, and the calibrationstandards included sixteen polystyrenes of precisely known molecularweight, ranging from a molecular weight of 500 to a molecular weight of5.2 million, and a polyethylene standard, NBS 1475. 10. The molecularweight distribution of the plastomers produced in this invention aretermed “narrow” that is to say an M_(w)/M_(n) less than 3, preferablyless than or equal to 2.5. The MI of the polymers of the invention aregenerally in the range of 0.01 dg/min to 200 dg/min, preferably 0.1dg/min to 100 dg/min, more preferably 0.2 to 50 dg/min and mostpreferably less than 10 dg/min. Contemplated densities of component A ofthe invention are in the range of 0.85 to 0.93 g/cm³, preferably 0.87 to0.92 g/cm³, more preferably 0.88 to 0.91 g/cm³.

[0064] The invention can be especially concerned with copolymerizationreactions involving the polymerization of one or more of the monomers,for example alpha-olefin monomers of ethylene, propylene, butene-1,pentene-1 1,4-methylpentene-1, hexene-1, octene-1, decene-1 and cyclicolefins such as styrene. Other monomers can include polar vinyl, dienes,norbornene, acetylene and aldehyde monomers.

[0065] In the case of elastomers, the polymer which may be producedinclude terpolymers of an ethylene-a-olefin-EODE(Ethylene-alpha-Olefin-Diene Elastomer) of high M_(w) and greater than0.3 weight % diene content, preferably greater than 2.0 weight % dienecontent. These polymers may be largely amorphous and have a low or zeroheat of fusion. As used herein the term “EODE” encompasses elastomericpolymers comprised of ethylene, an a-olefin, and one or morenon-conjugated diene monomers. The non-conjugated diene monomer can be astraight chain, branched chain or cyclic hydrocarbon diene having from 6to 15 carbon atoms. Examples of suitable non-conjugated dienes arestraight chain acyclic dienes such as 1,4-hexadiene and 1,6-octadiene;branched chain acyclic dienes such as 5-methyl-1,4-hexadiene;3,7-dimethyl-1,6-octadiene; 3,7-dimethyl-1,7-octadiene and mixed isomersof dihydromyricene and dihydroocinene; single ring alicyclic dienes suchas 1,4-cyclohexadiene; and 1,5-cyclododecadiene; and multi-ringalicyclic fused and bridged ring dienes such as tetrahydroindene, methyltetrahydroindene, dicyclopentadiene; bicyclo-1,5(2,2,1)-hepta-2,5-diene; alkenyl, alkylidene, cycloalkylidenenorbornenes such as 5-methylene-2-norbornene (MNB);5-propenyl-2-norbornene, 5-isopropylidene-2-norbornene,5-(4-cyclopentenyl)-2-norbornene, 5-cyclohexylidene-2-norbornene,5-vinyl-2-norbornene and norbornadiene.

[0066] Of the dienes typically used to prepare EPDMs, the particularlypreferred dienes are, 1,4-hexadiene (HD), 5-ethylidene-2-norbornene(ENB), 5-vinylidene-2-norbornene (VNB), 5-methylene-2-norbornene (MNB),and dicyclopentadiene (DCPD). The especially preferred dienes are5-ethylidene-2-norbornene (ENB) and 1,4-hexadiene (HD). The preferredEOD elastomers may contain 20 up to 90 weight % ethylene, morepreferably 30 to 85 weight % ethylene, most preferably 35 to 80 weight %ethylene. The alpha-olefin suitable for use in the preparation ofelastomers with ethylene and dienes are preferably propylene, 1-butene,1-pentene, 1-hexene, 1-octene and 1-dodecene. The alpha-olefin isgenerally incorporated into the EODE polymer at 10 to 80 weight %, morepreferably at 20 to 65 weight %. The non-conjugated dienes are generallyincorporated into the EODE at 0.5 to 20 to 35 weight %; more preferablyat 1 to 15 weight %, and most preferably at 2 to 12 weight %. Ifdesired, more than one diene may be incorporated simultaneously, forexample HD and ENB, with total diene incorporation within the limitsspecified above.

[0067] The elastomers may also be devoid of a diene and be a copolymerof two monomer types. Such copolymers may be elastomers of high M_(w),low crystallinity, and low ash. The copolymers may beethylene-alpha-olefin copolymers (EPC) of high M_(w). As used herein theterm “EPC” means a copolymer of ethylene and an alpha-olefin, notnecessarily propylene, which exhibits the properties of an elastomer.The alpha-olefins suitable for use in the preparation of elastomers withethylene are preferably C₃-C₁₀ alpha-olefins. Illustrative non-limitingexamples of such a-olefins are propylene, 1-butene, 1-pentene, 1-hexene,1-octene and 1-dodecene. If desired, more than one alpha-olefin may beincorporated. The EPC elastomers may contain 20 up to 90 weight %ethylene, more preferably 30 to 85 weight % ethylene, and mostpreferably 35 to 80 weight % ethylene.

[0068] In the case of polymers derived predominantly from propylenederived units, the polymers have the following features as a result ofthe presence of isotactic polypropylene sequences in the chain:

[0069] In one embodiment, a copolymer of propylene and at least onecomonomer, the comonomer being ethylene or an alpha-olefin. Comonomersinclude ethylene and linear or branched C₄ to C₃₀ alpha-olefins, orcombinations thereof. Preferred linear alpha-olefins include ethyleneand C₄ to C₈ alpha-olefins, more preferably ethylene, 1-butene,1-hexene, and 1-octene, even more preferably ethylene or 1-butene.Preferred branched alpha-olefins include 4-methyl-1-pentene,3-methyl-1-pentene, and 3,5,5-trimethyl-1-hexene. The propylenecopolymer of the SPC is preferably a random copolymer, as the term isdefined herein below.

[0070] The polypropylene copolymer has a crystallinity of from 2% to65%. Within this range of crystallinity, alternative lower limits ofcrystallinity can be 5% or 10%, and alternative upper limits ofcrystallinity can be 50%, 45% or 40%.

[0071] The crystallinity of the polypropylene copolymer of the SPC isderived from isotactic (or alternatively syndiotactic) polypropylenesequences in the copolymer. The amount of propylene in the SPC can befrom 65% to 95% by weight. Within this range, alternative lower limitsof propylene content in the SPC can be 70% or 80% by weight, andalternative upper limits of propylene content can be 92.5%, 90%, or 89%by weight.

[0072] The semi-crystalline polypropylene copolymer necessarily has anon-zero heat of fusion, due to the measurable crystallinity. Thecrystallinity can be calculated from the heat of fusion, using apreferred value of 189 J/g for 100% crystallinity and a linearrelationship between heat of fusion and crystallinity; see, B.Wunderlich, “Macromolecular Physics,” vol. 3, Academic Press (1980),esp. Chapter 8.4.2.

[0073] The polypropylene copolymer of the SPC preferably has a singlebroad melting transition. Typically, a sample of the polypropylenecopolymer will show secondary melting peaks or shoulders adjacent to theprincipal peak, and this combination is considered together as singlemelting point, i.e., a single broad melting transition. The highest ofthese peaks is considered the melting point. The polypropylene copolymerpreferably has a melting point of from 25° C. to 110° C. Within thisrange, alternative lower limits of the melting point can be 30° C. or35° C., and alternative upper limits of the melting point can be 105° C.or 90° C.

[0074] The weight average molecular weight of the polypropylenecopolymer can be from 10,000 to 5,000,000 g/mol, preferably 80,000 to500,000. The MWD (M_(w)/M_(n)) is preferably above 2. The MWD(M_(w)/M_(n)) may be less than 40, more preferably less than 5 and mostpreferably less than 3. In another embodiment, it is preferred that thepolypropylene copolymer has a ML (1+4)@125° C. less than 100, morepreferably less than 75, even more preferably less than 60, mostpreferably less than 30.

[0075] The polypropylene copolymer of the present invention preferablyis a random, crystallizable copolymer having a narrow compositionaldistribution. The intermolecular composition distribution of the polymeris determined by thermal fractionation in a solvent. A typical solventis a saturated hydrocarbon such as hexane or heptane. The thermalfractionation procedure is described below. Typically, approximately 75%by weight and more preferably 85% by weight of the polymer is isolatedas one or two adjacent, soluble fractions, with the balance of thepolymer in immediately preceding or succeeding fractions. Each of thesefractions has a composition (wt. % ethylene content) with a differenceof no greater than 20% (relative) and more preferably no greater than10% (relative) from the average weight % ethylene content of thepolypropylene copolymer. For purposes of the present disclosure, thepolypropylene copolymer is considered to have a “narrow” compositionaldistribution if it meets the fractionation test outlined above.

[0076] The length and distribution of stereoregular propylene sequencesin preferred polypropylene copolymers is consistent with substantiallyrandom statistical copolymerization. It is well known that sequencelength and distribution are related to the copolymerization reactivityratios. As used herein, the term “substantially random” means acopolymer for which the product of the reactivity ratios is generally 2or less. In contrast, in stereoblock structures, the average length ofPP sequences is greater than that of substantially random copolymerswith a similar composition. Prior art polymers with stereoblockstructure have a distribution of PP sequences consistent with these“blocky” structures rather than a random, substantially statisticaldistribution.

[0077] The reactivity ratios and sequence distribution of the polymermay be determined by C-13 NMR, which locates the ethylene residues inrelation to the neighboring propylene residues. To produce acrystallizable copolymer with the required randomness and narrowcomposition distribution, it is desirable to use: (1) a single-sitedcatalyst; and (2) a well-mixed, continuous flow, stirred tankpolymerization reactor which allows only a single polymerizationenvironment for substantially all of the polymer chains of preferredpolypropylene copolymers.

[0078] As general guidance when the molecular weight of the polymers istoo low, liquid phase separation in the manner described herein may behindered or made inefficient as an excessive amount of polymer mightthen be carried over in the lean phase. The precise boundary depends onsolvent composition and polymer composition as well as molecular weight.A rapid pressure let-down, generally greater than 20 bar/second,preferably 30 bar/second or more, more preferably 40 bar/second or more,even more preferably 50 bar/second or more, assists in inducingdisengagement of the two phases. This rapid pressure decrease preferablystarts from a pressure above the binodal boundary or LSCT and stops at apressure below the spinodal boundary. The preferred phase separation isby spinodal decomposition and is called pressure induced phaseseparation (PIPS). Also the liquid phase separator should provide asufficient residence time to permit the settlement of the lean andconcentrated phase at the lower end of the separator.

[0079] In the second aspect of the invention, molecular weight controlis exercised through control of hydrogen levels, which may besupplementary to control of molecular weight by control of thepolymerization temperature. In the second aspect the lean phase ispassed in liquid form to a means for removing hydrogen added to orgenerated during polymerization, which means comprises a means forcontacting a stripping vapor with the lean phase in a countercurrentflow arrangement to concentrate the hydrogen in the vapour phase forremoval from the lean phase recycle.

[0080] The stripping vapor preferably consists of a volatile monomersuch as ethylene (25). The means may include a stripping vessel (26) toremove hydrogen from the recovered solvent stream for use as thepolymerization feed (2). The stripping vapor advantageously has a lowhydrogen content, preferably below 5 mppm. The stripping vapor may beselected to be more volatile than other monomer or solvent components,be substantially devoid of contaminants that are deleterious to thepolymerization catalysts, be recoverable in the plant recovery system,and preferably be available at high enough supply pressure forintroduction into the stripping vessel (26) without the aid of separateadditional compression.

[0081] This aspect of the invention is especially applicable to plantlay-outs where reactors are arranged to operate in series and where theupstream reactor is operated under no or low hydrogen conditions toprovide a higher molecular weight fraction and where hydrogen is addedto a downstream reactor to provide a lower molecular weight fraction.

EXAMPLE

[0082] With reference to FIG. 1 the plant is arranged as follows:

[0083] Polymerization and Initial Separation of Polymer and Solvent

[0084] A feed for polymerization is passed through conduit (2) by acentrifugal pump (3). The feed contains A) hexane as solvent, B)monomer, generally the predominant monomer is ethylene or propylene, andC) comonomer which may be any copolymerizable alpha-olefin, and D) adiene or other polyene or cyclic copolymerizable material. The feed ispassed through a chiller or cooler (6) in which the feed is optionallychilled to a low temperature for subsequent adiabatic polymerization inthe two continuous stirred tank reactors (8) which are operated inseries (for simplicity, only one reactor is depicted in FIG. 1).Activator and metallocene catalyst may be premixed and added at (5)and/or (7) to one or both reactors (8). A scavenger, generally in theform of an alkyl aluminum such as tri-isobutyl aluminum or tri-n-octylaluminum is next added at (4) to minimize the impact of poisons in thefeed and in the reactor on the catalyst activity.

[0085] To complement the molecular weight control provided bycontrolling the polymerization temperature, hydrogen may be added to oneor both reactors through conduits (not shown).

[0086] The solution, containing polymer, which emerges from the reactors(8) through a conduit (11), is first treated with a catalyst killer,preferably water, added at (10) in a molecular solution in hexanesolvent to terminate the polymerization reaction. A heat exchanger (12)is arranged as part of a heat integrating arrangement and heated by alean phase emerging from an upper layer (20) in a liquid phase separator(14), and provides an initial increase in the temperature of the polymersolution in the conduit (11). A trim heat exchanger (16), operating byusing steam, hot oil or other high temperature fluid, further increasesthe temperature to a level suitable for liquid phase separation. Thesolution then passes through a let down valve (18) where a pressure dropis created which causes the separation of the polymer solution andsettlement into the lean phase (20) and a polymer rich phase (22) belowit.

[0087] It is important to note that no energy consuming pump is requiredto provide a pressure increase in the conduit (11) between the reactors(8) and the separator (14) as the polymer containing solution ispropelled by the pressure from the pump (3).

[0088] Treatment of Lean Phase

[0089] The lean phase (20), after being cooled by the heat exchanger(12), aforementioned, is cooled further by a cooling device (24), passedthrough a surge tank (26) adapted for stripping out the hydrogen andthen submitted to in-line chemical analysis at (43) to determine theconcentration of monomer and comonomer in the solvent. This cooled leanphase (43) is combined with fresh feed of solvent and monomer (30) toprovide the desired concentrations and then passed through a drier (32)which serves to remove any unreacted water used as the catalyst killeror present in the fresh feed supplied or any impurity in the recycledsolvent and monomer as will be explained.

[0090] The surge tank (26) is arranged in the form a vessel (26)suitable for stripping out hydrogen by means of ethylene as a strippingvapor as is shown in FIG. 3. The lean phase issuing from the cooler (24)is passed through a conduit (27) to a liquid distributor arrangement(100) located inside the vessel (26) in an overhead space in an upperpart thereof. The liquid distributor consists of a performated pipedistributor with holes (102) on the bottom. The distributor sprays thelean phase downward inside the vessel (26). Lean phase collects in thelower part of the vessel (26). Part of the ethylene to be added to thefeed conduit (2) is supplied as stripping vapor through line (25). Thestripping vapor is supplied to a vapor sparger arrangement (104) locatedinside the vessel (26) submerged in the lean phase collected in thelower part of the vessel. The vapor sparger consists of multiple disksof microporous media (103) arranged on a plurality of rings, arrangedconcentrically. Vapor bubbles rise from the vapor sparger arrangement(104) through the liquid to the surface into the overhead space. Thevapor in the overhead space is passed through conduit (108) for furthertreatment as described below. The liquid in the lower part is passedthrough conduit (110) for treatment as will be described below.

[0091] In the vessel (26) a countercurrent flow of the liquid feed (27)and the stripping vapor (25) occurs. At the stage where the bubbles ofstripping vapor rise through the liquid, the ethylene in the vapor isdissolved in the liquid and hydrogen in the liquid is taken up by thebubbles. Hence the liquid issuing through conduit 110 is enriched byethylene which can be subjected to polymerization when recycled. A firstequilibrium stage can so be approximated. In the vessel (26) vaporspace, the rising vapor extracts more hydrogen from the atomizeddroplets issuing from the nozzles (102) so that a second equilibriumstage can be approximated. The vapor issuing through conduit 108 thuscontains a large proportion of the hydrogen contained in the liquidintroduced though nozzles 102. Substantially 2 equilibrium stages ofseparation can be achieved in a single flash vessel. Over 90%, sometimesover 97%, of the hydrogen present in the lean phase can be removed inthis way.

[0092] The stripping vapor supplied is ethylene, which is a volatilemonomer indigenous to the process. Its use minimizes additionaloperating costs and raw material consumption.

[0093] The vapor from conduit (108) is routed to the reflux drum (39) oftower (36). Partly it is processed to recover valuable components,principally volatile monomers such as ethylene and propylene, byfractionating tower (36) and its overhead vapor compression/condensationsystem for recycling through conduit (43) to the inlet side of the drier(32). The part mainly comprising hydrogen and any other non-condensablesmay be flared at (112).

[0094] A less preferred alternative is for part of the lean phaserecycle to be flashed in a single stage flash vessel without theaddition of stripping vapor. This, however, only permits limitedhydrogen removal and detracts from the benefit of recycling the leanphase in its liquid state without energy intensive evaporationprocesses.

[0095] In single reactor and in series reactor arrangements usingmetallocene catalysts systems varying amounts of hydrogen may beproduced by beta-hydride abstraction, even when no hydrogen is injectedinto the reactor. The amount may vary with metallocene selected. Itsmolecular weight reducing effect may be accommodated by appropriateselection of the reactor operating temperature. A substantial amount ofthis hydrogen may remain unreacted in the reactor effluent stream (11).Reducing the amount of hydrogen recycled in this stream in the mannerdescribed above may be is advantageous to permit adjustment of themolecular weight independent of the polymerization operating temperatureby removal of the generated hydrogen or by addition of hydrogen from anexternal source, generally in the feed conduit (2).

[0096] In series reactor operation as described herein, the ability toremove hydrogen can be exploited advantageously to widen the molecularweight split between the reactors and to broaden the molecular weightdistribution beyond what would otherwise be possible. The feed suppliedto the upstream reactor can have a hydrogen content below that whichwould prevail if hydrogen generated by beta hydride elimination remainedin the recycle. Additional extraneous hydrogen can be added to thedownstream reactor to provide a hydrogen content above that which wouldremain if hydrogen from beta hydride elimination were to remain in therecycle.

[0097] Effective removal of the hydrogen thus provides a facility whichenables the range of bimodal compositions produced in series reactor layouts to be increased. It also permits the selection of a broader rangeof metallocene catalyst systems regardless of their tendency to generatehydrogen through beta hydride elimination or their sensitivity to thepresence of hydrogen in the polymerization mixture.

[0098] Treatment of Polymer Rich Phase

[0099] The concentrated polymer rich phase is passed to a low-pressureseparator (34) where evaporated solvent and monomer are separated fromthe more concentrated polymer solution emerging from the liquid phaseseparator (14).

[0100] The evaporated solvent and monomer phase is passed throughconduit (35) in a vapor phase to the purification tower (36) operatingby distillation to separate a light fraction of the highly volatilesolvent and unreacted ethylene and propylene on the one hand and heavierless volatile components such as hexane and any toluene used to dissolvecatalyst or activator and unreacted diene type comonomers on the otherhand. Use of toluene can be reduced under appropriate circumstances by asuitable selection of catalyst components and catalyst preparationconditions such as increases in catalyst solution temperature toincrease the solubility of the catalyst components to reach a pointwhere so little toluene is present that no separate process for theremoval of the toluene are required.

[0101] A gear pump (38) conveys the by now even more concentratedpolymer to a vacuum devolatilizing extruder or mixer (40), where again avapor phase is drawn off for purification, condensed and then pumped toa purification tower (50). A heavy fraction of toluene used as catalystsolvent and diene such as ethylene norbornadiene (ENB) comonomer oroctene-1 comonomer are recovered by this purification tower (50). TheENB or octene can be recycled through outlet (54). Alternative heavycomonomers, such as ENB and octene, may thereby be stored in separatestorage vessels (55, 56), which facilitates rapid product transitionsbetween different product families (e.g. EP(D)M and EO plastomers, whilestill enabling eventual recovery of the valuable unreacted comonomers.This capability further enhances the flexibility of this process toproduce a wide variety of dissimilar products.

[0102] The polymer melt emerging from (40) can then be pelletized in anunderwater pelletizer, fed with water chilled at (42), washed and spundried at (44) to form pellets suitable for bagging or baling at (46).

[0103] Polymerization of Differing Polymers

[0104] The operation of the plant can be best described with referenceto the Table 1 on the following page. This takes as examplespolymerization processes to make a low molecular weight plastomer (asdescribed generally above); a higher molecular weight elastomer (asdescribed above) and a high propylene content ethylene copolymerpolymerized as described above. TABLE 1 Process Conditions of thePlant/Process of the Invention in Varying Operating Modes PolymerPolymer Amount of Solution Solution Polymer Devolatilized PolymerizationUpstream Let- Downstream Lean Polymer Rich Polymer from Feed IntoReactor Inside Reactor Down Valve Let-Down Valve Phase Phase ExtruderPlastomer 0 or down to − 150-200° C.; 220° C.; 40 or 220° C; 40 bar;220° C.; 40 220° C.; 40 bar; High 15° C.; 120 bar 100 or 120 bar; 100bar; 15-22 15-22 wt % bar; <0.1 wt 30-40 wt % total; 50 bar 15-22 wt %wt % polymer polymer % polymer polymer monomer partial polymer pressure.Elastomer 0 or down to − 100° C; 100 or 220° C.; 100 bar; 220° C.; 40bar; 220° C.; 40 220° C.; 40 bar; Medium 15° C.; 120 bar 120 bar; 8-128-12 wt % 8-12 wt % bar; <0.1 wt 30 wt % total; 50 bar wt % polymerpolymer polymer % polymer polymer monomer partial pressure. Predominant0 or down to − 50 or 60° C.; 200° C.; 100 bar; 200° C.; 40 bar; 200° C.;40 220° C.; 40 bar; Low Propylene 15° C.; 120 bar 120 bar; 7-8 7-8 wt %7-8 wt % bar; <0.1 wt 30-35 wt % Content total; 50 bar wt % polymerpolymer polymer % polymer polymer Copolymer monomer partial pressure.

[0105] To make plastomer in FIG. 1, the feed temperature is reduced bythe chiller (6) to 0° C. Aluminum alkyl is added as scavenger in amountsappropriate to the poison content of the feed. Alternatively the processof WO9722635 (Turner et al.) incorporated herein for US purposes. Thepressure is raised by the centrifugal pump to 120 bar. The feedcomprising largely solvent and up to 50 bar partial pressure of ethyleneand butene or hexene or octene comonomer then enters the first of thetwo series reactors (8). Catalyst and activator is added to the reactors8 in amounts to create the desired polymerization temperature which inturn is related to the desired molecular weight. The heat ofpolymerization increases the temperature to 150 to 200° C. to form aplastomer without the use of hydrogen (although H₂ may be used). At theoutlet of the second series reactor, the polymer concentration is in therange of from 15-22 wt %. The general conditions may be as described inWO 99/45041 incorporated herein for US purposes.

[0106] Water is then supplied at 10 to kill the polymerization reactionwhich might otherwise continue in the presence of surviving catalyst,unreacted monomer, and elevated temperature.

[0107] The heat exchanger (12) raises the temperature initially and thenthe further heat exchanger (16) causes a further temperature rise to220° C. A rapid pressure drop results as the polymerization mixturepasses through the let-down valve (18) into the liquid phase separator,with the pressure dropping quickly from 100 Bar to 40 bar. The pressuredifferential between that at the outlet of the pump (3) and the outletof the let down valve 18 is solely responsible for causing the feed andthe polymerization mixture to flow through the reactor (8) and theconduit (11) including the heat exchangers (12) and (16).

[0108] The details of liquid phase separation by passing through a lowercritical solution temperature (LCST) boundary is explained withreference to FIG. 2. Polymerization takes place at 100 or 120 bar in thepolymerization reactor(s) at the pressure also prevailing upstream ofthe pressure letdown device at a level as shown by line A. Thetemperature is maintained and/or raised to a range marked by the bracketshown at B to between 150 and 200 or 220° C. At the prevailingtemperature, the pressure is dropped along the arrow to a level markedX. As the temperature is dropped across the letdown valve from 100 barto 40 bar, the polymerization mixture passes from a homogeneous singlephase, through the lower critical solution temperature boundary markedLCST, to a two-phase (L-L) region. (i.e. for a given temperature, thepressure starts at a pressure above the highest of thepressure-temperature curves representing the upper critical solutiontemperature (UCST), the LCST, and the vapor pressure, and the pressureafter the let-down for the given temperature is below thepressure-temperature curve representing the spinodal boundary and abovethe pressure-temperature curve representing the vapor pressure) Thepressure drop is sufficiently fast to avoid formation of a continuouspolymer and to form a discontinuous solvent/monomer phase. The pressuredrop across the region bounded by the LCST (binodal) boundary and thespinodal boundary must be especially rapid to induce phase separation byspinodal decomposition, which leads to rapid phase separation andsettling.

[0109] Level X is above another phase boundary marked Vapor pressurebelow which the mixture enters a V-L-L region in which it is part vapor,and part two phase liquid. The pressure at level X at the exit of theseparator is sufficiently high so that no vapor is formed.

[0110] Inside the separator (14) an upper lean phase is formed with lessthan 0.1 wt % of polymer and a lower polymer rich phase with 30 to 40 wt% of polymer. The concentration is approximately double to triple thatof the polymerization mixture fed to the separator (14). After furtherremoval of solvent and monomer in the low-pressure separator (34) andthe extruder (40), polymer can be removed from the plant containing lessthan 1 wt %, preferably with 0.3 wt % or less, even more preferably <0.1wt % of volatiles, including water.

[0111] If the use of the plant is now compared with the row in Table 1marked elastomer, it can be seen that although the polymerizationtemperature is lower than for plastomer, and the polymer concentrationemerging from the reactor is lower (its viscosity will be similar tothat for plastomers), the same separation process and plant can be usedto give an output which is somewhat lower (reflecting the reducedefficiency of the polymerization process at lower temperatures). Withtwo reactors in series, the disclosure of WO 99/45047 (Harrington etal.) may be used, which document is incorporated herein for purposes ofUS law. Generally speaking, in a series lay out it is preferable thatthe first reactor operates at temperatures between 0 to 110° C. and thesecond reactor operates between 40 to 140° C. Preferably the firstreactor operates at temperatures between 10 to 90° C. and the secondreactor operates between 50 to 120° C. Most preferably, the firstreactor operates at temperatures between 20 to 70° C. and the secondreactor operates between 60 to 110° C. With appropriate control ofprocess conditions and poison levels temperature of this order ofmagnitude can also be obtained where one reactor only is used or tworeactors are used under the same process conditions.

[0112] The same can be said about the row in Table 1 marked “Predominantpropylene content copolymer” where the temperature is lowered to allowthe less reactive propylene monomer to form a sufficiently highmolecular weight. The general conditions described in WO 00/01745, whichare fully incorporated herein by reference for purposes of US patentpractice, can be used. In the runs, the polymerization temperaturevaried between 28 and 70° C.

[0113] While the process windows have been illustrated using prior artdisclosures which suggest metallocene selection and the suitableoperating window for a given polymer type, to the extent that priorpublished patent specifications are used to assist in such illustration,it should be kept in mind that these patent specifications did notprovide the separation and recycle and purification conditions in acontinuous plant with a recycle permitting full exploitation of theproduct capabilities of high activity metallocene catalyst systems atwhich the invention is best operated. The invention provides a plant andprocess which uses the given metallocene catalyst systems disclosed tomake the target polymers at high metallocene activity, under a widerange of polymerization conditions and with considerable energy andinvestment savings.

[0114] Advantages

[0115] It can be seen that the plant and process illustrated above in anon-limiting manner, permit polymerization and subsequent polymerseparation across a broad range of temperatures to yield polymers ofwidely varying average molecular weights and comonomer contents withcatalyst optimized for operation at low or high operating temperatures.A process plant according to the invention is capable of production ofplastomers, elastomers, and predominant propylene content copolymers bychanging only the substituents of the reaction mixture and the processconditions.

[0116] The plant has a low energy consumption because the extent of feedrefrigeration and re-pressurizing by pumping can be greatly reduced.Furthermore, no heat of vaporization is required to separate the polymerrich and polymer lean phases in the separator (14), and the heat in thelean phase is efficiently used to increase the temperature of thepolymerization mixture entering the separator (14). Little solvent andmonomer has to be purged to atmosphere. The evaporated materialsrecovered from the low pressure separator (34) can be fractionated toallow its direct purification without intervening condensation in afractionating tower and at the same time assisting the flow of the finalpolymer/solvent mixture before extrusion into the extruder to minimizepumping requirements. Variation in the pressure of the low-pressureseparator (34) may be used to control viscosity of the polymer/solventmixture into the devolatizer (40), thereby extending the range ofpolymers that can be processed to those with very high molecular weight.

[0117] Widely varying molecular weights and molecular weightdistribution can be obtained using the hydrogen stripping arrangementwhich can be simply integrated and requires no additional extraneousmaterials or the evaporation of large volumes of recycled lean phaseliquid.

[0118] All documents cited herein are fully incorporated by referencefor all jurisdictions in which such incorporation is permitted and tothe extent they are not inconsistent with this specification. Alldocuments to shich priority is claimed are fully incorporated byreference for all jurisdictions in which such incorporation ispermitted. Although dependent claims have single dependencies inaccordance with U.S. practice, each of the features in any of thedependent claims can be combined with each of the features of one ormore of the other dependent claims dependent upon the same independentclaim or claims.

1. A process for continuous solution polymerization of a feed ofolefinically unsaturated monomer in a hydrocarbon solvent underpressure, having a continuous stirred tank reactor arrangement to whicha single site catalyst is supplied, to form a polymer containingpolymerization reaction mixture, and downstream thereof separating meansfor continuous separation of the solvent and unreacted monomer from themixture, which separating means includes at least an initial liquidphase separator to separate the polymerization mixture into a lean phaseand a concentrated phase wherein A) a high capacity, low viscosity pumpraises the pressure of the feed to at least 75 bar and causes themixture to pass from the reactor arrangement through a heating stage upto a pressure reducing means upstream of the liquid phase separatorthrough the action of the pump and in the absence of further pumpingmeans between the reactor arrangement and the pressure reducing means;and B) a catalyst killer is added downstream of reactor arrangement andupstream of the liquid phase separator to suppress furtherpolymerization in the separator of the heated polymerization mixtureundergoing separation, the lean phase being passed through a cooler andoptionally a drier back to the inlet side of the pump; the concentratedphase being subjected to additional solvent removal downstream to obtaina solid polymer:
 2. Continuous solution polymerization process accordingto claim 1 in which a low pressure separator receiving the concentratedphase from the liquid phase separator operates at a pressure of at least2 bar gauge at a level sufficient to convey the volatile phase removedfrom the concentrated phase to a fractionating tower for purificationwithout further means of compression.
 3. Continuous solutionpolymerization process according to claim 1 or claim 2 in which thesolvent is a low boiling alkane based solvent and the polymer in thepolymerization mixture contains at least 50 mol % ethylene Or propylenederived units out of the total monomers present.
 4. Continuous solutionpolymerization process according to claim 1 or claim 2 in which thepolymerization mixture from the reactor arrangement is heated to thetemperature for separation in the separator successively by exchangersand the lean phase from the separator is used to supply heat to theupstream one of said heat exchangers
 5. Continuous solutionpolymerization process according to any of the preceding claims in whichthe lean phase is passed in liquid form to a means for removing hydrogenadded to or generated during polymerization, which means comprises ameans for contacting a stripping vapor with the lean phase in acountercurrent flow arrangement to concentrate the hydrogen in thevapour phase for removal from the lean phase recycle.
 6. A plant foroperating a continuous solution polymerization process of a feed ofolefinically unsaturated monomer in a hydrocarbon solvent underpressure, which plant contains a continuous stirred tank reactorarrangement to which a single site catalyst may be supplied to form apolymer containing polymerization reaction mixture, and downstream ofthe reactor arrangement separating means for continuous separation ofthe solvent and unreacted monomer from the mixture, which separatingmeans includes at least an initial liquid phase separator to separatethe polymerization mixture into a concentrated phase and a lean phase,wherein A) a high capacity, low viscosity pump is arranged for raisingthe pressure of the feed to at least 75 bar and a heating arrangement isprovided, to thereby cause the mixture to pass from the reactorarrangement through the heating arrangement up to a pressure reducingmeans upstream of the liquid phase separator (14), no additional pumpingmeans being provided between the reactor and the pressure reducingmeans; and B) means are provided for introducing a catalyst killerdownstream of reactor arrangement and upstream of the liquid phaseseparator to prevent further polymerization in the separator of theheated polymerization mixture undergoing separation, and means areprovided for cooling the lean phase; and an optional drier, for removalof catalyst killing residues may be arranged to convey a dried recycledstream back to the inlet side of the pump
 7. A process for continuoussolution polymerization of a feed of olefinically unsaturated monomer ina hydrocarbon solvent under pressure, having a continuous stirred tankreactor arrangement, to which a single site catalyst is supplied, toform a polymer containing polymerization reaction mixture, anddownstream thereof separating means for continuous separation of thesolvent and unreacted monomer from the mixture, which separating meansincludes at least an initial liquid phase separator to separate thepolymerization mixture into a lean phase and a concentrated phase,wherein A) a pump maintains a pressure of at least 75 bar in thepolymerization reactor and the temperature in polymerization reactorarrangement is maintained lower than or the same as that in theseparating means and a heating arrangement is provided upstream of theseparating means to optioially heat the reaction mixture to raise thetemperature in the separating means so that, upon a pressure drop in theseparating means, the mixture passes through the lower critical solutionphase boundary regardless of the initial polymerization temperature; andB) a catalyst killer is added downstream of reactor arrangement andupstream of the liquid phase separator to suppress furtherpolymerization in the separator of the heated polymerization mixtureundergoing separation, the lean phase being passed through a cooler andoptionally a drier back to the inlet side of the pump; the concentratedphase being subjected to additional solvent removal downstream to obtaina solid polymer.
 8. Continuous solution polymerization process accordingto claim 7 in which the liquid phase separator and a low pressureseparator receiving the concentrated phase from the liquid phaseseparator operate at a pressure of at least 2 bar gauge, at a levelsufficient to convey the volatile phase removed from the concentratedphase to a fractionating tower for purification.
 9. Continuous solutionpolymerization process according to claim 7 or 8 in which the solvent isa low boiling alkane based solvent and the polymer in the polymerizationmixture contains at least 50 mol % ethylene or propylene derived unitsout of the total monomers present.
 10. Continuous solutionpolymerization process according to claim 7 or 8 in which thepolymerization mixture from the reactor arrangement is heated to thetemperature for separation in the separator successively by exchangersand the lean phase from the separator is used to supply heat to theupstream one of said heat exchangers
 11. Continuous solutionpolymerization process according to any of the preceding claims 7 to 10in which the lean phase is passed in liquid form to a means for removinghydrogen added to or generated during polymerization, which meanscomprises a means for contacting a stripping vapor with the lean phasein a countercurrent flow arrangement to concentrate the hydrogen in thevapour phase for removal from the lean phase recycle.
 12. A processplant suitable for continuous circulation of a solution polymerizationreaction mixture, said plant comprising: a) a pressure source, b) apolymerization reactor, downstream of said pressure source, c) apressure let-down device, downstream of said polymerization reactor,said let-down device capable of reducing the pressure of the reactionmixture at a selected temperature pressure-temperature curve to apressure below the spinodal boundary pressure temperature curve at arate of at least 20 bar/second, and d) a separator, downstream of saidpressure let-down device, wherein said pressure source is sufficient toprovide pressure to said reaction mixture during operation of saidprocess plant to produce a single-phase liquid reaction mixture in saidreactor and a two-phase liquid-liquid reaction mixture in said separatorin the absence of an additional pressure source between said reactor andsaid separator.
 13. The process plant of claim 12 wherein said plant iscapable of production of plastomers, elastomers, or predominantpropylene content copolymers by changing only the substituents of thereaction mixture and the process conditions.
 14. A process forcontinuous solution polymerization of a feed of olefinically unsaturatedmonomer in a hydrocarbon solvent under pressure, having a continuousstirred tank reactor arrangement, to which a single site catalyst issupplied, to form a polymer containing polymerization reaction mixture,and downstream thereof separating means for continuous separation of thesolvent and unreacted monomer from the mixture, which separating meansincludes at least an initial liquid phase separator to separate thepolymerization mixture into a lean phase for passing through a recycleback to the reactor arrangement and a concentrated phase, wherein thelean phase is passed in liquid form to a means for removing hydrogenadded to or generated during polymerization, which means comprises ameans for contacting a stripping vapor with the lean phase in acountercurrent flow arrangement to concentrate the hydrogen in thevapour phase for removal from the lean phase recycle.
 15. A processaccording to claim 15 wherein the stripping vapor comprises ethylene foruse in subsequent polymerization.
 16. A process according to claim 15 orclaim 16 wherein a pair of reactors are arranged in series and hydrogenis added to the downstream reactor to produce a lower molecular weightfraction.
 17. A plant for continuous solution polymerization of a feedof olefinically unsaturated monomer in a hydrocarbon solvent underpressure, having a continuous stirred tank reactor arrangement, withinlets for feed and for a single sited catalyst components, to form apolymer containing polymerization reaction mixture, and downstreamthereof separating means for continuous separation of the solvent andunreacted monomer from the mixture, which separating means includes atleast an initial liquid phase separator to separate the polymerizationmixture into a lean phase for passing through a recycle back to to thereactor arrangement and a concentrated phase, a conduit being providedfor passing the lean phase ed in liquid form to a stripping arrangementfor removing hydrogen added to or generated during polymerization, whichstripping arrangement comprises a means for contacting a stripping vaporwith the lean phase in a countercurrent flow arrangement to concentratethe hydrogen in the vapour phase for removal from the lean phaserecycle.